Selective conversion of unstable liquids



SELECTIVE CONVERSION oF UNSTABLE LIQUIDs Filed Nov. 19, 1962 UnitedStates Patent O 3,239,0349 SELECTIVE CNVERSIN F UNSTABLE LQUIDS RichardG. Graven, North Castle, and Vernon 0. Bowies,

Bedford Township, Westchester County, NPY., assignors to Socony MobilOil Company, Inc., a corporation of New York Filed Nov. 19, 1962, Ser.No. 233,690 21 Claims. (Cl. 208-143) The present invention relates to aprocess for the selective conversion into more stable substances ofunstable liquids which tend to deposit solids upon heating. It isparticularly concerned with improved control of vaporization ofintermediate liquid products between reaction stages in selectiveconversion processes having both liquid or mixed phase reactions andvapor phase reactions. In a preferred embodiment it involves theselective hydrogenation in two or more stages of a liquid hydrocarbonmixture containing aromatic hydrocarbons, olefins, diolelins and sulfurcompounds, and especially a mixture of narrow boiling range.

Selective hydrogenation serves many purposes and the instant inventionis particularly concerned with a process in which an unstablehydrocarbon mixture of the type mentioned above is hydrogenated in atleast two stages of increasing severity to prepare a stable product fromwhich valuable aromatic hydrocarbons can readily `be separated bysolvent extraction with a solvent such as diethylene glycol. In suchextraction, it is relatively easy to separate benzene and other aromaticcompounds `from paraiiins or naphthenes; but this is not true ofseparating benzene from unsaturated aliphatic components and especiallyfrom organic sulfur compounds in the mixture.

To prepare a suitable feed for the solvent extraction, it is necessaryto convert the organic sulfur compounds to a readily separable materialsuch as hydrogen sulfide gas, to saturate the unstable gum-formingdioletins and also to saturate the mono-oleiins without convertingaromatic hydrocarbons into naphthenes by excessive hydrogenation.Although it is easy to specify the reactions with hydrogen for obtainingthese results, performing them in commercial practice has -been anentirely different matter. There is an increasing demand for theproduction of aromatic hydrocarbons from petroleum so that the suppliesof these compounds are not restricted to the current production level ofthe steel and coking industries. Despite this demand there was still nofully satisfactory commercial method for the hydrogenation of suchmixtures of aromatic and unsaturated aliphatic hydrocarbons prior to thepresent invention.

It is not feasible to saturate and desulfurize such feed stocks in asingle operation because the relatively high temperatures suitable forhydrodesulfurization also promote the formation of coke and oletinspolymers or gums and may induce the hydrogenation of aromatics tonaphthenes. Prior to the present invention, conducting the hydrogenationreactions in two stages to avoid or minimize the aforesaid deficiencieshas not been entirely satisfactory by reason of the accumulation ofpolymeric deposits that reduce the activity of hydrogenation catalysts,thereby necessitating frequent regeneration. Such deposits also plug uppiping and other equipment. Not only thermal polymerization but alsocatalytic polymerization 4must be minimized as some hydrogenation anddesulfurization catalysts also catalyze the polymerization of diolens.While various techniques are known for at least partially reducingypolymer formation in hydrocarbous at elevated temperatures,nevertheless polymer formation has remained a critical problem incommercial rrr ICC

plants for the selective hydrogenation of charging stocks oi the typedescribed.

An object to the invention is to provide an improved process for theselective conversion of unstable liquids into relatively stable liquidshaving little or no tendency to polymerize or otherwise deposit solidsupon standing.

Another object of the invention is to provide an improved process thatincludes simple, direct and positive control of the vaporization ofliquid at a heat labile stage during a selective conversion process.

A further object of the invention is to provide an improved selectiveconversion process for the direct regulation of the vaporization of aliquid of narrow boiling range.

An object of the present invention is to provide an improved process forthe selective hydrogenation of an unstable mixture of organic compounds.

Another object of the invention is to provide a process forhydrogenating a mixture boiling below aboutA 500 F. of aromatic andunsaturated aliphatic hydrocarbons without the formation of substantialquantities of napht'hene's or polymers.

A further object of the invention is to provide an improved process forthe selective, nondestructive hydrogenation of unstable,thermally-cracked, petroleum products with a boiling range below about500 F., which hydrogenation is performed in stages of increasingseverity without excessive deactivation of the contact catalysts.

Still another object of the invention is to provide an improved processfor 4the selective, nondestructive hydrogenation of a mixture ofaromatic and oleiinic hydrocarbons boiling 'below about 5007 F., andpreferably below about 275, in which contact catalysts are kept onstream for substantially longer periods.

A still further object of the invention is to provide an improvedprocess for the selective hydrogenation of an unstable hydrocarbonmixture in which not more than a very small amount of polymeric materialis formed and it is removed from the intermediate product stream priorto the final hydrogenation reaction.

Other objects and advantages of the invention will be apparent to thoseskilled in the art -upon consideration of the following detaileddisclosure in which all temperatures are expressed in terms of degreesFahrenheit, all proportions in terms of weight and all boiling point orranges of temperatures are measured at atmospheric pressure according tothe ASTM procedure unless otherwise stated.

The present invention relates to the selective conversion of unstableliquids with a pronounced tendency to deposit solids upon heating by aprocess which includes partially converting unstable compounds in aliquid feed into more stable substances within a confined initialreaction zone under conversion conditions in which a considerableproportion of said feed is maintained in the liquid phase, vaporizing asubstantial portion of the eiiluent liquid from the initial conversionreaction by controlled heating, and passing the gaseous phase derivedfrom said initial eiiiuent through a confined conversion zone whilesubstantially completing the conversion of unstable components of saidfeed; and in which the improvement comprises separating in an enlargedseparation zone a liquid iiux in an amount equal to at least about 0.5%(at least 5% being preferred) of said liquid feed from said gaseousphase, withdrawing said liquid ux at a substantially constant rate froma pool thereof maintained in said separation zone and regulating saidcontrolled heating operation in direct response to the rate ofcollecting said liquid flux in said separation zone as determined fromthe level of said pool.

Narrower aspects of the invention include, inter alia the use of contactcatalysts in the various reaction and conversion zones, increasedseverity of conversion conditions in a subsequent conversion stage incomparison with the initial reaction, retaining specified proportions ofthe initial eliluent in the liquid state after the controlled heatingand separation, heating the efliuent during passage through a restrictedtransfer conduit (a term employed herein in its broad sense .to includeheaters, etc.), heating the initial effluent by the injection of ahotter stable gaseous reactant either upstream or downstream from theseparation-zone or both, which heating by injection may be Vregulated asdescribed rhereinbefore, and the sources, amounts and circulation ofliquid linx. Feeds of narrow boiling range (eg. those with a breadthofboiling range which does not exceed about 150 or does not exceed about80 on the rFahrenheit scale) are particularly contemplated since thevaporization control of-Y such materials by conventional methods lacksprecision;-

whereas the present method is especially suitabler for'the vaporizationof liquids of very narrow boiling ranges or even those having boilingpoint rather than a range because regulation of the controlled heatingis independentY of the boiling and vaporization characteristics of theini-v within certain specified ranges to providea hydrogenation effluentin which at least about 35%, and preferably at least about 60%, of thediolens have been at least partially saturated and in which an amountequal to at least about 20%, and preferably at least about 60%, of theliquid feed remains in the liquid phase. The, Bromine Number of thenormally liquid fraction of said eiiluent is also desirably reduced atleast 25% below that of the liquid feed.` Said conditions which areregulated therel include maintaining a hydrogen partialpressure withinthe range of Vabout 200e800 (about 300-600 being pre-l ferred) pounds'per'square inch absolute pressure( herein-A after abbreviatedp.s.i.a.), an hourly spacevelocity'with-IV in the range of about0.2-15.0, and preferably about 0.5-8.0, based on the Volume of liquidfeed, the hydrogen temperature within the broad range of about 75-300F.,- and preferably :not over about 250. Controlled Vaporization ofliquid in said hydrogenation eiiuent and separation of theV gaseous,phase thereof in an enlarged separationV zone is eifected in vthepresence of a liquid liuX, for example an amount of ilux equal to atleast t about 5% (preferablyover 10%) of said liquid feed.

Y gaseous material derived from saidfyaporization step, is

passed through a subsequent conversion zone in contact with a poroussolid conversion catalyst ofat least mod-'.-

erate hydrogenation activity and high` desulfurizration activity at asubstantially higher average temperature than in said initialzonewhilelregulating conversion conditions within certain specified rangestoproduce agconversion eiliuent with a normally liquid fraction having aBromine Number'less than about 4, .preferably below about 2, and anorganic sulfur content `belowabout 20 p.p.m., and -Y preferably below 15p.p.m. The speciiiedpranges of conversion conditions includemaintaininga hydrogen par-vv tial pressure with-in the range of about `20G-800(about 1 30G-600 being preferred) p.s.i.a., an hourly space velocitywithin the range of about 0.2-6.0 (about,0.54.0 being; preferred) basedon the volume of original liquid feed,

the total hydrogencharge (unreacted hydrogen plus any newly introducedhydrogen) within the' range of about G-10,000 (preferablyfabout2,000-5,000) s.c.f.b. ofV said liquid feedand an inlettemperaturewithinv the` z' wide range of about S50-700, and preferablyk about e'In' performing the preferred modificationof/ the instant process, afeed stock with a pronounced tendency toward' jundesired polymerizationis subjected to a selective hydrogenation pro-cess in which it is firsthydrogenated mildly inthe liquid -or mixed phase; the resultingefliuent'is vaporized under controlled conditions and iinally treated ina the gaseous phase with hydrogen under more severe conditions. Theinitial hydrogenation isiconducted at a temf perature sufficiently lowto avoid or minimize boththermal and catalytic polymerizationgwhilehydrogenating a substantial portion and usually almost all of the-diole- `tins, including all of the more reactive ones. During the' rinitial hydrogenation little, if any, desulfurization isV acf.

complished and a substantialproportion of the mono-olelins usuallyremain unsaturated here as, under some con-V without depositingpolymeric solids on the equipment, if

the vaporization is accomplished by heating in a carefully controlledmanner up until the liquid and gaseous, phases are separated from oneanother completely and rapidly in an enlarged separation zone. It is tobe noted that the Y charge within the range of about SOO-6000, andpreferv ably about 1200-3000, standard cubic feet per barrel f therspecial-precautions to a temperature suitable for de- (hereinafterabbreviated s.c.f.b.) of liquid and a feed Such vaporization isproducedby a controlled heating.

operation that is regulated in direct response to the rate of collectingsaid flux in said separation zone and preferably in direct response tothe rate of collecting the liquid fraction of fresh initialhydrogenation eiiiuent. The flux may be any liquid substance or mixturewhich is rniscible v with and unreactive with said liquid feed includingimmediate reaction products, recycled partially or fully hydrogenatedproducts of the instant process or vextra neous liquids, preferablyhaving a substantial content of hydrocarbon vapor phase is sharplyV andcompletely separated for `the tir-st time from the circulating fluxliquidand 4the, unvaporized fraction of the Veffluent' by removal of thevapor phase and lnot'by evaporating therliquidkf.

phase to dryness orl a close approach to dryness.

This gaseous phaseis further heated withoutanyI fursulfurization andalso for olen saturation and then itis subjected to la catalytic'conversion with hydrogenat a distinctly higher temperature than in theinitial hydro-` genat'ion. In this conversion,-thesaturationrofall-remaining mono-olefins and diolens is substantially corn- Vpletedandorganic sulfur compounds are converted into.

hydrogen sulfide without any appreciablevpolymerfor.

mation occurring in either the preliminary heating or in, f the vaporphase reaction even though the catalyst is euse.

tomarily of a type of high polymerization,potentialsY When the catalystin the initial reactor is in relatively, .Y

fresh condition, most of the hydrogenation of monoole fins (often morethan as well fas diolens, occurs there.k S little? hydrogenationr of thehydrocarbonsVoc-V tinued use, more of lthe hydrogenationl loadlrisshifted to Y Y V; the second or linal reactorY andsubstantial'increasesbe#A tween theV inlet :and outlet temperatures'. of this reactor); arethen apparent., Y Y Y Eventually as theY activity As the startingmaterial, any mixture of aromatic and unsaturated aliphatic hydrocarbonswith organic sulfur compounds may be employed in this embodiment processif the final boiling point of the liquid does not exceed about 500. Anarrow boiling range material, for example, one having a boiling rangebetween about 140 and 275, is desirable, and preferably a -chargingstock boiling in the range about 160 to 220.

As `a source of aromatic hydrocarbons, the liquid feed -desirablycontains a total -of between about 20 and 90% aromatics, especiallybenzene and toluene. Typically, it also has substantial contents ofdiolefins and olefins as evidenced by Diene Numbers of about to 22 whichmeasure the proportion of lconjugated diolefins as determined by thema'leic anhydride condensation method and Bromine Numbers of about to 30which represent the total content of unsaturated aliphatic hydrocarbons.Feeds with Diene and Bromine Numbers as high fas about 40 and about 75respectively may also be processed according to this phase of thepresent invention. The -organic sulfur content is typically about to 300p.p.m. :and may be as high las about 700 p.p.m.

In other utilizations of the instant process, the charging stock neednot be rich in aromatic hydrocarbons. For instance, in producing astable gasoline blending stock from a pyrolysis liquid, a feedcontaining 6 to 20% aromatics compounds is typical.

Feed stocks of the nature described are unstable as they tend to formpolymeric gurus readily. It has been found desirable to keep the periodof storing them as brief as possible in order to minimize theintroduction of gum into the present process. In addition, it isrecommended that the liquid feed stock be free of dissolved oxygen andbe ,stored in the substantial absence of oxygen or air, for example,under a blanket of an inert gas such as nitrogen. This prolongs theactivity of the catalysts usable in this process. Such feed stocks :aregenerally obtainable by severely thermally cracking a petroleum fractionsuitable for the manufacture of gasoline or light `olefins, as exempliedby ethylene. It is preferred here to depentanize the cracked product. Apartially suitable feed is one within an end point not exceeding 220 anda maximum gum content of less than 15 milligrams per 100 milliliters.

The total consumption of hydrogen in this modification of the inventionVaries of course -with the particular feed stock employed; but ingeneral, it is in the range of about 150-800 s.c.f.b. of liquid feedstock. A typical value is 300 s.c.f.b. with a charging stock of Dieneand Bromine Numbers of l5 and 24 respectively, and the consumption isusually found to be less than 500 s.c.f.b. Substantial ,excesses ofhydrogen have been specified -hereinbefore to avoid a drop in thehydrogenation rates as a result of an inadequate supply of hydrogen.Although pure hydrogen may be used, it is customarily supplied as amixture of hydrogen `and gaseous hydrocarbons in the off gases of unitsfor reforming naphthas or hydrodesul-furizing gas oils, etc. The gascharge preferably has 1a hydrogen content of at least 60% by volume butgaseous mixtures with as little as 40% hydrogen may be used.

The partial pressure of hydrogen in the t-wo or more reactors -isimportant in avoiding undesired side reactions, such as the formation ofgum or coke on the catalysts. It should be maintained within the rangeof about 200- `800 p.s.i.a., in which the 300-600 p.s.i.a. range is-preferred. The total pressure in the reactors is not critical, but it-s'hould not be so high as to interfere 'significantly with thevaporization of the `feed and reaction produ-cts described herein. Amajor proportion of the product gases with much unconsumed hydrogen isrecycled to the process after `any excessive quantities of hydrogensulfide have been scrubbed out and this usually constitutes ra majorproportion of the total quantity of gases charged to the reactors.

The charging stream of combined recycle and makeup gases containinghydrogen is divided into several streams.

A substantial quantity of hydrogen must be introduced into the firstreaction zone along with the liquid feed, and unreacted `hydrogen ispresent in the effluent of that reaction which is subjected to furtherhydrogenation reactions. While in theory all of the hydrogen-rich gasrequired for the series of selective hydrogenations can be charged tothe initial reactor, this is not particularly desirable in practice.Especially since the circulating gas m-ay be employed after heating to ahigh temperature in a furnace as a heat source to aid in vaporizing theeffluent from the initial reactor and also to regulate the temperatureof the vapor phase charge of hydrocarbons and hydrogen to a subsequentreactor. Alone this hydrogen-rich gas stream can be heated withoutdecomposition or other difficulty to a temperature several hundreddegrees higher than it is possible to heat the liquid or mixed phaseeffluent of ythe first reactor without the coincident deposition ofpolymeric gum or coke. Such deposit-ion from the mixed phases can occurat temperatures of 300 and even lower. In serving as a heat source, asubstantial part of the total circulating gas, say about 30 to 85%, isheated to a temperature in the range of about 500 to 950, and preferablyin the range of about 60G-850, while the unheated balance of the gas ischarged to the initial reactor. One stream desirably containing morethan half of this heated gas is used to supply the final temperatureincrement to the initial mixed phase effluent just prior to entering theenlarged separating and vaporizing chamber, and the remainder may beintroduced into the wholly gaseous stream leaving the top of saidchamber on its way to the second stage conversion reactor as the finalheat increment to adjust the charge to the desired inlet temperature.

In the preferred modification of the novel process, a catalyst of highhydrogenation activity is required for the initial reaction zone as itmust hydrogenate at a relatively low temperature the more reactiveconjugated diolens and usually at least some of the other olefins, butits polymerization activity must be relatively low in order to avoid theformation of gums which will deactivate the catalyst. While suitablehydrogenation catalysts also incidentally possess relatively highdesulfur-ization activity initially, this property drops ofi rapidly ina period of a few days to a week, because such catalysts are readilypoisoned with respect to desulfurizing ability at desulfurizingtemperatures by feeds containing much organic sulfur.

These qualities of the catalysts may be defined in terms of arbitraryactivity indexes which are described herein. Unless otherwise stated,all such indexes are measured using fresh new catalyst. The activityindexes enable one to clearly differentiate between the two or morecatalysts employed at various stages in the instant process.

For delineating hydrogenating `activity in a preferred embodiment, twodifferent indexes are available. The hydrogenation activity index isdefined herein as the percentage or proportion of isoprene which isconverted to pentenes and pentanes when a blend of 8-l0% isoprene and50-500 p.p.m. of thiophene sulfur in benzene is passed over the catalystwith 1500-3000 s.c.f.b. of hydrogen gas at F., 300 pounds per squareinch gage (hereinafter designated p.s.i.g.) as the total pressure and aliquid hourly space velocity of 5. Thus a conversion of half of theisoprene present, or 4.5% out of a total of 9.0% isoprene present,signifies that the activity index is 50. For the initial catalyst, ahydrogenation activity index of at least about 40 is recommended.

In determining the benzene conversion index as another and usuallysupplemental measure of hydrogenation activity, a sulfur-free mixture of17% benzene and 83% cyclohexane is passed through the catalyst undertest at 400 F. and 400 p.s.i.g. with 1500-3000 s.c.f.b. hydrogencirculation and a liquid hourly volumetric space velocity of 2. The testmixture must be sulfur-free inasmuch as organic sulfur in a content assmall as 50 p.p.m., and even less in the form of hydrogen sulfide,totally inhibits the hydrogenation of benzene with such catalysts underthe specified reaction conditions. A suitable catalyst for the initialreactor has a benzene conversion index of at least about 50, meaningthat half of the benzene present or 8.5% is converted into cyclohexane,but an index of Vabout 100 is typical with the preferred catalysts.

Another means for designating suitable catalysts for the first reactoris the polymerization activity index. This is another arbitrary indexwhich equals the `percentage of isoprene that is polymerized when 25 cc.of a mixture of 8-l0% isoprene in benzene is heated with 5 cc. of t-hemonomer due to polymerization is calculated 4by difference between theisoprene content of the reaction product -and that of the test blendcharged. A satisfactory catalyst for the first reactor has apolymerization activity index less than about 35, as polymerizationthere is undesirable.

An arbitrary desulfurization activity index may also be used in thismodification of the present invention, princi pally for determiningsuitable catalysts for the subsequent desulfurization operation. Thisindex is the percent reduction in sulfur content obtained when a blendof pure compounds consisting of 10% hexene and 10% isoprene in 80 volumepercent of benzene with a total thiophene sulfur content of 500 p.p.m.Vis passed over the catalyst in question at 500 F. and 450 p.s.i.g.together with between 1500 and 4000 s.c.f.b. of hydrogen at a liquidhourly volumetric space velocity of 2. For sui-table desulfurization thefinal stage catalyst desirably has a desulfurization activity index ofat least about 80, both fresh and t after one week of operating with thetest feed s-tock.` It has been observed that good catalysts for theinitial hydrogenation zone also have desulfurization indexes in the;

80-100 range initially but these catalysts are quickly poisoned ordeactivated by a sulfur content equivalent to that in the test feed sothat after one week of operation the activity index -is in the 0 to 50range.

Catalysts of substantial acid activity are not desirable for thisparticular process since they produce unwanted cracking reactions, sosilica-alumina catalyst supports are usually avoided. However, While itis preferable that 'the catalyst lsupport is lsubstantially free ofhalogens, a relatively low halogen content up t0 about 0.5% may betolerated. Furthermore, a Acatalyst is favored which is substantiallydevoid of alkylation activity and thus does not promotethe alkylation ofaromatics with olefins.

A variety of catalysts of differing chemical constitution may beemployed in the initial hydrogenation step as long `as they have thenecessary Iactivity described herein. Platinium in amounts ranging fromabout 0.105 to 2.0%, preferably about 0.\2,1.0%, supported on variousaluminas, and especially gamma and chi alumina, is suitable as 4are theother noble' materials in group VIII of the Periodic Table of Elements,such as rhodium land palladium. The concentration of palladium in suchcatalysts may be about 0.05-10% and about 0.2-2.0% is preferred for thepurpose. NickelV either unsupported or on known supporting materials inconcentrations ranging down to about 10% nickel in the compositecatalyst Ialso provides satisfactory results, as does copper chromite.For instance, good hydrogenating characteristics for the yfirst reactorare obtained with 55% nickel supported on kieselguhr. genation activityat low temperature, palladium or platinum on gamma alumina arerecommended, palladium being preferred for the initial reaction, sinceits greater By reason of their high hydro.

activity catalyzes the desiredhydrogenation.reactions at a temperatureabout 100 lower than in the case of platinum. The palladium composite isdesirably Vpromoted in sorne instances with a quantity of chromia inlthe same range as the palladium. Among the many suitable speciliccatalysts are 5% palladium on `activated'carbon and 0.6% platinum onetaalumina of less than 0.01% chlorine content. The manufacture of suchcatalysts is well known in the art and accordingly is not describedhere.

The catalyst or catalysts employed inthe second or a final reactoroperate under quite different reaction con. ditions from Ithose in theinitial reactor. The charge is enti-rely inithe kgaseous phase yandsubstantially higher.

temperatures are required in the second reactor to desul.

furize the intermediate product; and =to complete the.

hydrogenation of the less reactive unsaturated hydrocarbons, namely the:mono-oleiins and any remaining dioleiins. This catalyst may also beVdefined in terms of an arbitrary desulfurizationlactivity.indexaotivity asset forth hereinbefore in the -1100 range,which isfretained for a period of at least one week and usually. muc-hlonger;

It should have at leastmodera-te hydrogenation activity. Y The increasedtemperature of -the fnalreaction greatlyV increases the actualhydrogenation activity of these catalysts. Also it hasbeen foundvthatfsome and perhaps The preferred catalyst Ifor theyfnal, stage is asulfded.

composite of cob-alt molybdate on the surface of gamma alumina.

In illustration, the ypresulfidingfo-r finalstep in thepreparation of apreferred type of`desulfuriz-ation catalyst may desirably be performedin situ inthe reactor.

Afresh contact catalyst containing cobalt molybdate on` the surface of asuitablek support such as gamma alumina or a catalyst regenerated totheYoxide state byjcombustion with air diluted by steam isgsubjectedfirst-to prereduction for six hours at*v 700 p.-s.i.g.` and 700ol with ahydrogen-rich recycle gas substantially free of hydrogen sulfide.Following ,the .prereduction step athe catalyst is then contacted with acirculating stream .of mixed hydrogen sulfide and hydrogen underconditions such that the minimum partial pressures are 8 p.s.i.a.y forhydrogen zsuliideand 'p.s.i.a. in the casebfhydrogen and the temperatureis in the range of 500 to 700. ment is concluded at a temperature; ofabout 700 after being continued until the sulfur content of the contentof the composite -catalyst rises to the range. 6.5-7.5%

whereupon the catalyst is ready to be placed on-strearn'.,V

Later during desulfurizing operations, the sulfur in the catalyst dropsfrom that range to an equilibrium content;

of about4.6%

Returning now tothe first reactor, suitable ranges of reactionconditions for this preferred processhave been described earlier and theactual reaction conditions are selected and regulatedvwitliin thoseranges in amanner' known to `those ,skilled inlthe art to produce aninitial hydrogenation eiuent in which at lea-strabout 35 andpreferablyr'at least .60%, of the original diolefins have been converted,into mono-olefinsor p-arafn's andin whichV an yamouutequal to at leastabout`.20%, and prefeffect of one operating Ivariable upon another arewell understood by those skilled in the art and need not be explained indetail here. For instance, if the Vdegree ofy The treatv hydrogenationtends to drop Ibelow the minimum specified or perhaps below thepreferred value, this condition can be corrected by increasing the feedtemperature or decreasing ythe space velocity or both. Also if theproportion of initial reactor effluent in the liquid phase drops belowthe minimum specified, the feed temperature may be decreased, the spacevelocity increased to reduce the total exothermic lheat generated andprovide a greater quantity of reactants to absorb the heat liberated, ort-he pressure increased or any combination of these measures may beemployed in reducing the degree of vaporization in the initial reactor.Using the same circulating gas, an increase in total pressures of courseresults in a corresponding increase in hydrogen partial pressure.

With a fresh catalyst, either nevi or regenerated, it is obviously mosteconomical to maintain the feed or charging temperature at the lowesttemperature at which the gaseous and liquid components of the charge arereadily available thus avoiding any heating or cooling expense. The feedtemperature should of course be within the stated range and, in thepreferred operation, the feed temperature is maintained at asubstantially constant value within the narrow range of 75-l90 F., anddesirably in the lower part of that range, While the catalyst is fresh.Generally, this temperature is subsequently increased either graduallyor by steps but not beyond about 300 F. in order to maintain a diolefinsaturation above the specified minimum, and preferably substantiallyconstant degree of saturation, as the hydrogenation activity of thecatalyst decreases with continued use. Even with a fresh catalyst, thefirst hydrogen treatment customarily does not fully saturate all of theolefnic or unsaturated aliphatic compounds for the Bromine Numberreduction usually is in the range of about 25-95%.

Regeneration of the initial hydrogenation zone catalyst is required whenthe Diene Number reduction is less than the prescribed minimum of 35%,or the degree of vaporization exceeds 80%, or both, even after the feedtemperature has been adjusted upward to the stated maximum. These arebetter criteria than prescribing a maximum outlet temperature for theinitial reactor inasmuch as the degree of vaporization of-the efiiuentand the degree of saturation of its more reactive original components,are more significant than the outlet temperature in the instantembodiment of the novel process. In addition, it appears that themaximum permissible outlet temperature can vary considerably fordifferent feed stocks over the range of about 275 to 400. For instance,a reactor outlet temperature of 325 is considered excessive for certainlow boiling feed stocks but will give satisfactory results with otherfeeds boiling at higher temperature ranging up to end points near 500'.

An attempt to define acceptable outlet temperatures is not feasible inconsideration of the variations in permissible pressures in the reactor.For example, an outlet temperature of 325 would be suitable for arelatively high reaction pressure in retaining the necessary proportionof liquid phase efiiuent while a lower temperature would be necessary ifthe minimum pressure were employed while all other conditions were heldconstant. Thus as a result of the close interrelation of the variousoperating conditions, it is more significant to define the initialhydrogenation in terms of the regulation of certain reaction conditionswithin restricted ranges to provide an intermediate product in which acertain proportion is retained in the liquid phase and a certain amountof the more reactive feed components are at least partially saturated.

An entirely different situation prevails at the outlet of thedesulfurization reaction zone as it is unlikely that any exothermcreated by the stated reaction conditions can reach a temperaturesufficiently high to deactivate the catalyst. To permit the use ofordinary construction materials, the maximum outlet temperature shouldnot exceed about 850 Although catalysts in the form of palladium orplatinum supported on alumina retain their activity for extremely longperiods, as for instance, 3 months or more in the Case of palladiumcatalysts, regeneration of the catalyst is eventually necessary and thismay be readily accomplished by heating the reactor to a temperature ofabout 700-900 for a palladium-alumina bed while passing a gas containingl or 2% oxygen therethrough. A diluent is usually introduced with theair to avoid excessive regeneration temperatures which can reducecatalyst activity considerably. Nitrogen or flue gas may be usedgenerally for that purpose and the more convenient medium of steam maybe utilized as the diluent with a palladium catalyst.

The desulfurization or final stage catalyst is conventionallyregenerated in similar fashion at even longer intervals of about 6months or more. In the case of a sulfided composite of cobalt andmolybdenum on alumina, this converts the cobalt and molybdenum compoundsto oxides and a presulfiding treatment such as the one describedhereinbefore is employed to restore the catalyst to its original form.

It has also been found that purging the initial catalyst with hydrogenat 200-500 p.s.i.a. and 750-850" for 16-4 hours sometimes serves toregenerate certain catalysts, such as palladium, almost as effectivelyas conventional regeneration by combustion with air diluted to an oxygencontent of 1 or 2 percent. Accordingly, it is contemplated that, in theabsence of severe deactivation of the catalyst, this catalyst may beregenerated several times by such treatment with hydrogen-rich gasbefore it is necessary to regenerate the contact agent by the combustiontechnique.

Only a limited amount of hydrogen sulfide may be tolerated withoutsubstantial deactivation by certain of the catalyst suitable for thefirst reaction stage of this species of the invention. Although thisloss of activity may be readily restored either by regeneration of thecatalyst in the usual fashion or the hot hydrogen treatment describedearlier, frequent regenerations reduce the over-all efficiency of theprocess. Accordingly, in the case of a platinum catalyst supported onalumina, it is desirable that the concentration of hydrogen sulfide inthe gaseous phase should not exceed 0.05 p.s.i.a. and preferably shouldbe less than 0.03 p.s.i.a. The effect on a palladium catalyst issimilar. Organic sulfur generally has a lesser effect on the catalystand it is a relatively simple matter to control the hydrogen sulfidewhich is introduced in the hydrogen-containing gas by simply passingeither or both of the make-up and recycle gases through an alkalinescrubber, or other unit for removing hydrogen sulfide such as adiethylamine absorber.

Under severe conversion conditions, for example a high desulfurizationtemperature in combination with a low space velocity of perhaps lessthan 1, a sulfided composite of cobalt and molybdenum on alumina maycatalyze the hydrogenation of a part of the aromatic hydrocarbons, asexemplified by the conversion of benzene to cyclohexane. This is usuallyundesirable and may be easily avoided by inhibiting the reaction bymaintaining a concentration of sulfur compounds in the charge equivalentin inhibiting effect to at least about 50 p.p.m. of thiophene sulfur(e.g., about 20 p.p.m. of hydrogen sulfide). Where the charge containsless of such compounds it is a simple matter to supply additionalhydrogen sulfide in the hydrogen-rich gas which is introduced upstreamof the final reactor. Selection of a make-up gas of suitable hydrogensulfde content or by-passing the recycle gas around the caustic sodascrubber are some of the methods useful in attaining any additionalinhibiting effect.

Despite the unstable nature of the hydrocarbon feed stock, especiallywhen subjected to substantially complete vaporization, very little ifany gum is formed in the first reactor of this particular process. Therelatively low reaction temperature is not conducive to thermalpolymerization. A catalyst having little or no polymerization activityis employed. A substantial proportion of the reaction mixture ismaintained in the liquid phase to avoid approaching the point of drynessin the reactor. In addition, the usually substantial aromatic content ofthis liquid'makes it a good solvent for polymeric gums, so the thoughthis technique has been suggested in the prior art.

Such procedure deposits polymer either in the heater or in the catalystmass or both, and stoppages of this nature call for much cleaning and/orregeneration that reduce the overall operating efiiciency. Instead theinstant process is concerned with vaporization of the initial eiuent inthe presence of a flux liquid. This may be accomplished by variousmethods, one of which involves a combination of stages in which theinitial hydrogenation effluent is gradually heated undei goodtemperature control in the presence of a flux, preferably circulating insubstantial quantity through the transfer line between the` initialreactor and a vaporizing and separating chamber of enlarged crosssection. In that chamber vaporization of the initial original feed andproducts thereof is completed to the desired extent of about 90 to 99%and seldom more than 99.5%. The small but significant balance ofunvaporized effluent is withdrawn from the process as a liquid leavingthe enlarged chamber and carrying a small amount of polymerformedrduring the vaporization operation and possibly also in theinitial hydrogenation step or perhaps present in the original chargestock. Once this separation of the gaseous and liquid phases has beenaccomplished, there is no longer a tendency toward any significantpolymerization in the gaseous phase containing the major proportion ofthe hydrocarbons even when it is heated up to temperatures of 350 to 700which would have produced an unacceptable degree of polymerization inthe mixed phase material from the initial reactor. Y

The gradual heating of the initial effluent to effect controlledvaporization during passage of the eluent through the restrictedtransfer conduit (including heater passages, etc.) leading from theinitial reactor to the vaporizing and separating chamber may beaccomplished by several means. One, comprises an optional but preferredtechnique in which a circulating liquid flux at a substantially highertemperature, than the reffluent typically of the order of 75-200 higher,is injected into the initial hydrogenation effluent near the outlet ofthe first reactor. It will be appreciated that the exotherm of theinitial reaction has already increased the temperature of -this effluentsubstantially above the temperature of the feed to that reactor. Thetemperature of the mixture of flux and reaction effluent is preferablyincreased further during passage through an indirect heater which isdesirably heated with steam or another easily controllable medium foreven heating. A relatively low'temperature difference between theheating and the heated media is highly desirable to provide the gentleheating that minimizes polymerization in such equipment. Indirect heatexchange is recommended for the major heat input into the stream passingthrough the transfer conduit. In one embodiment of the invention, anamount of the initial reaction efluent equal to betweenabout and 50% 0fthe original liquid feed may be retained in the liquid phase in thiscontrolled heating operation whenifurther Vaporization is' subsequentlyproduced by the injection into the heated effluent of a stable gaseousreactant at a higher temperature as described hereinafter. Finally, andpreferably closely adjacent to the inlet of the separator,

Accomplishing 12 an additional stream of the hydrogen-rich gas used inthis process may be injected into the mixture at a temperature severalhundred degrees higher than-the temperature of the mixture.Vv Thisdirectcontact heating with jet of hot gases is an optional but highly'desirable feature which minimizes polymer' deposition on equipmentsurfaces. With each of these increments of heat,l more of the lirstreactor elluent is converted in the transfer conduit from the liquidphase into the gaseous -state under conditions in which the presence atallrtimes of a substantial liquid phase assistsin preventing or at leastin `minimizing the deposition of polymeric material or heated.

surfaces. The enlarged crosssectionl of the chamber provides goodconditions for separating the two-phases bycatch any traces of entrainedliquid in the rising vapors.

The supply of steam to the indirect heater may be manually controlledvto maintain a predetermined temperature in the vaporizing chamber assteady as possible, but far better :results are usually obtainable inregulating the steam supply in response to the liquid level in `the.separating chamber.

In brief vsuch a regulating system involves. controlling the input ofsteam manually, but preferably automatically, in direct response tothesignals of a conventional liquid level indicator or: controller attachedto theV vaporizing and separating chamber. The removaly of liquidstreams from that chamber Ias well as any input ofV external flux isdesirably maintained at constant flow rates under the` regulation ofvautomatic flow controllers; therefore, a rise in the liquid level in theseparating chamber represents a decrease in the vaporization of the`initial hydrogenation.

effluent and a fall in that level-means=that the eliiuent is beingvaporized in a greater degree. To maintain a steady degree ofvvaporization more steam or less steam respectively is supplied to theindirectheater. The heating steam may be adjusted Iby means of avalve'in :the steam supply line or one in the line used for drainingcondensed heating steam from the heater.

Conventionally, control 'of vaporization of a generally similar natureis regulated in response tothe temperature of the vapor or perhaps theliquid temperature., Such control is subject to the usual deviationsencountered in efforts to obtain `precise elevated temperaturemeasurements thatarise from radiation or evaporization-of liquid on atemperature sensing element, etc. not particularly satisfactory forliquids of narrow boil-ing range, such as the preferred feedstof thepresent invention,l

inasmuch as a small temperaturey differential of only a few degrees at asubstantially elevated temperature generally is related to a largeVdiiferential in the proportion Vof liquid vapor-ized.: Thus control of fheating of the liquid in direct response tothe actualproportion ofunreacted feed Either manual or automatic eontrolof'the heating of 1 theinitial eflluent in direct -responseto the liquid level in the flashchamber `may also; be extended-to controlling the quantity of heatsupplied by the stream .of hot hydrogen-v rich gas injected into thetransfer line near the inlet of the vaporizer pot'.

line or on the temperature at the jcharge outlet of the furnaceldescribed hereinafter for heating that gas. Also,

it is possible to control both the heat input to the indirect heaterthrough which t-he initial Veffluent passes yand the heat furnished tothe effluent by the: hot hydrogen-rich stream in response to the .liquidlevel controller on the vaporizing andV separating chamber,

Moreover, it is This regulation may be `exercised eitherl on thequantity of said gas being admitted to the transfer Y However, it is;usually preferred from a standpoint of practical opera,

13 tions t-o apply such regulation only to the steam input to theindirect heater.

The liux liquid comprising the liquid fraction of the effluent from theinitial reactor and any inert liquid miscible therewith that isintroduced into the transfer line may perform several functions beforebeing separated from the gaseous portion of that eluent in theseparation chamber. It minimizes or inhibits gum format-ion at thiscritical stage of the preferred process wherein a stream of mixedgaseous and liquid lhydrocarbons containing gum-forming precursors iscarried to a relatively high degree of vaporization by heating, for theflux prevents the effluent from approaching dryness too closely, forexample, not closer than about based on the original liquid feed rate.Secondly, the circulating flux serves as an economical and relativelygentle direct heating medium for vaporizing a portion of the initialeffluent. Finally, t-he flux liquid prevents, or at least minimizes,t-he deposition of any gums or polymeric solids on the pipes and otherapparatus by reason of its washing -action on the surfaces thereof andits solvent characteristics which enable it t-o retain in solution anypolymeric material Whether formed at t-his stage or earlier.

Although any hydrocarbon liquid of suitable boiling and stabilitycharacteristics may be employed as the flux in this particularembodiment of the new process, it is preferred that the content ofaromatic compounds should amount to lat least 15% to improve itscapability for dissolving gummy material. A flux liquid from an externalsource may be used, and it is suggested that its volatility should besuficiently low that a major proportion and preferably substantially allof the iiux remains in the liquid state under the conditions in thevaporizing chamber while its resistance to coking and polymerizationshould desirably be at least as good as that of the initial efuent. Itsboil- -ing range is preferably located between about the boiling pointof benzene and about 950. However, an economical and readily availableux liquid may be obtained very simply by merely reducing the -heating inthe transfer line between the first reactor Iand vaporizer pot at thestart of a run in order to retain a larger than usual proportion of theinitial reactor etiluent in the liquid phase until a suicient body offlux liquid has been built up in the system. This of course amounts to:accumulating the least volatile fraction of the feed stock as theliquid liux.

The rate of recirculating the flux liquid in the preferred :process may`amount to at least 5%, and preferably at least 10%, of the rate ofintroducing the liquid feed stock into the frst reactor, and lesseramounts may be Irecirculated where an appreciable proportion of theinitial effluent is retained in the liquid-phase throughout thevaporizing step. As used herein, all flux (liquid efllue-nt plus anyadded liquid) quantities or rates relate to the proportions at themoment when the maximum degree of vaporization of the initial eiiluentis attained; and, of course, the proportion 'of material in the liquidphase reaches its minimum--namely, the instant of separation of thegaseous and liquid phases-rather than at the confluence of a circulatingtlux stream with the initial hydrogenation effluent. Much higher fluxcirculating rates can be employed ranging up to 40%, and even to 200% ormore, for the only lreal limitations are physical ones relating to thecapacities of the equipment and economic ones relating to pumping costsand the cost of larger equipment. When the total proportion of liquid inthe transfer line and heater is ample by a substantial margin to avoiddryness and bathe the walls of the equipment, furthe-r increases in theflux circulation rate do not achieve a corresponding or even asignificant reduction in the amount of polymer formed in the system oreven in the polymer concentration in the ux liquid; hence, circulationrates provide no important advantages.

The concentration of polymer in the circulating flux is dependent on thesmall but significant proportion of spent ilux withdrawn from thevaporizing step and from the instant process either intermittently orpreferably continuously. This spent liquid is derived from anunvaporized fraction of the eiuent of the initial reactor or from asupply of external flux or from both sources, and over any substantialperiods the rate of withdrawal must equal the supply from these sources.Under the preferred steady state conditions, reducing the degree Iofvaporization of the initial eflluent and correspondingly increasing thespent flux withdrawal results in a decrease in the polymer concentrationin the circulating ux and vice-versa. As indicated earlier, this removalof spent flux liquid amounts to at least about 0.5% and desirably about1 to 10% based on the liquid feed rate. While the amount may be larger,it is generally uneconomical to withdraw much more in the liquid phasefor purification or further processing. In actual practice of theinstant embodiment a llow controller on the spent flux line from thevaporizing chamber may be adjusted manually as needed to keep the gumcontent of the circulating liquid low enough to avoid the deposition ofpolymerio material in the equipment; for example, by keeping the gumcontent below about 200 milligrams per milliliters.

Where a flux liquid from an external source is supplied to the system ata constant and usually relatively low rate, it is possible to vaporize acorrespondingly greater proportion, in fact the whole, of the liquideiiiuent fraction yof the first reactor. However, it is preferable toretain the least volatile 0.5% or 1% of said effluent in the liquidstate in order to keep the temperature as low as possible during thevaporizing operation. For example, with all percentages based on `theliquid feed rate, one may continually charge 5% of a hydrocarbon -oilhaving an atmospheric boiling range of 600700 and a major proportion ofaromatic hydrocarbons to the separating chamber as circulating flux, andrecycle 25% liquid from this pot to the transfer line immediatelydownstream of the first reactor; then vaporization of the efflunt-fluxmixture in the transfer line may be controlled by appropriate heating toretain 1% of the initial efuent in the liquid phase in that chamber and6% spent llux may be withdrawn continually from the bottom thereof inmaintaining steady operations.

The size and shape of the separating and vaporizing chamber are notcritical. In avoiding or minimizing appreciable entrainment of liquiddroplets in the vaporous phase that is leaving, it is desirable to keepthe velocity of the gaseous phase relatively low, perhaps 2 feet/ secondor less. This can be achieved by providing a reasonably largecross-sectional area perpendicular to the direction of gas flow in theupper part of the vessel. On the other hand, where the heat forvaporization is regulated in response to liquid level in the chamber, itis desirable to have a relatively small cross-sectional area in theneighborhood of that level in order that a significant change in levelwill occur whenever a significant change in the degree of vaporizationof initial effluent occurs. Such factors pose no great problems, asthere is no necessity `for maintaining a constant crosssectional areathroughout the length tof the vessel. As one illustration, the vesselmay be in the general form of a double cylinder having a lower sectionof considerably smaller diameter than the upper section.

After separation of the flux liquid from gaseous material derived fromthe efiluent of the initial reactor, this gaseous phase is heated ifnecessary to bring its temperature up to the desired inlet temperaturelof the second reactor and its proportion of hydrogen is boosted, ifnecessary, to the desired level for that reactor by the introduction ofa hydrogen-rich gas. These steps may be combined, if so desired, byintroducing the extra hydrogen-containing gas at a substantially greatertemperature, say about 100 to 400 more, than that of the gaseous phaseleaving the separating chamber. This is one of the suitablemethods ofmaking the final temperature adjustment in the charge to the nalreactor. It is preferably accomplished by regulating the volume of fuelgas burning in a furnace for heating circulating gas and consequentlythe outlet temperature of that circulating gas stream either manually orautomatically in response to `signals from a temperature sensing devicelocated in the conduit leading to the inlet of the second reactor.

For a better understanding of the nature and objects of this invention,reference should be had Ito the detailed description and exampleshereinafter taken in conjunction with the accompanying drawing which isa simplied flow sheet or schematic representation of the process of thepresent invention. It will ybe appreciated that many details well knownto petroleum engineers have been omitted from the drawing orsimplifiedfor simplicity and greater clarity including Ipumps,- valves,alternate and parallel piping and equipment andV control equipment,especially instruments for indicating, recording or regulatingtemperature, pressure, level, flow, etc.

EXAMPLE 1 Turning now to the drawing, a freshly-distilled stream ofthermally cracked and depentanized gasoline (170- 220 B. R.) of thecomposition set forth in part in column 1 of Table I hereinafter entersthe feed conduit Z' at ambient temperature and a pressure of 740p.s.i.g.

for catalytically hydrodesulfurizing gas oil is admitted iny pipe 4 at apressure of 750 p.s.i.g. The Vquantity and` composition of this gas arespecified in column 2 of Table I. This make-up gas joins the recycle gasstream, whichy lis described later, in conduit 6. The resulting mixturehas a temperature of 125 F. and its composition and rate of flow aredesignated in column 3 of the table.

Half of the mixed gas stream in conduit 6 is taken oif in the valvedline 8 for purposes that willbe apparent later. The other half of thegaseous material continues to travel along pipe 6 until it joins thehydrolysis liquid hydrocarbons in conduit 2, and this gas-liquid mixtureof the composition and ilow rate given in column 4 of Table I passesthrough the heater 10 where its temperature is adjusted to 115 (hereindesignated as the feed temperature) by heating, if necessary, on its Wayto reactor 12. This charge temperature produces good results with thecatalyst described hereinbefore which has been partially deactivated bya liquid feed containing undesired polymeric materials.

Column 4 of Table I setsforth the total charge to the rst or initialreactor 12 which contains a xed or stationary catalytic bed ofchromia-promoted palladium on a gamma alumina support in the form of3/16 diameter i cylinders 5716 long. Based on the total weight, there isa surface deposit on the alumina of 0.50 percent of palladiurn metal and0.51 percent of chromium in the form of oxides.

The reaction conditions in the first reactor 12` are:

In the rst stage reaction the primary reactionisone of thenondestructive hydrogenation ofk diolefnsespecially conjugated diolens,accompanied by considerably less t saturation of the less reactivemono-olens. The temperat tures are below the levelrequi'red fordesulfurization and no significanthydrogenation vof aromatics orpolymerization takesplace there.

Any trace of gum formed in the catalyst bed dissolves in the descendingliquid and the reaction eluent is drawn off at the bottom of the reactorviaV conduit 14 .in'which it is transported to heater 15.' vA minorportion of the liquid feed stock or reaction products thereof vaporizesin reactor 12 'as a result ofthe heat evolved in the exothermichydrogenation lreaction.V

A circulating iiux liquid at 350 is injected from the conduitl16 intothe products in pipe 14 partly to increase the temperature ,of theinitial reactoreluentabout 45 thus promoting its vaporization but chieyyto reduce any tendency toward the deposition I.of any gummy 'solidsfinthe transfer` line 14. This flux liquid isJdraWn off near.

the bottom of the vaporizer pot 18in pipe 16 and recirculated by pump22.at the rate of 9,220,1bs./l1r.`0r720 b./d. This liquid is composed ofthehigherY boiling hy,-

drocarbons of the initial reactor eflluent' which'. are -retained in theliquid phase and a small quantity of dissolvedv polymeric material. Thelatter is a by-product of the present process and is readily soluble inthe benzene and other aromatic hydrocarbons constituting most of theliquid flux.

Two other modes of heating the rst reaction euent` are also employedduring its passage to `thevaporizervpotv 18.l Saturated steam at220-p.s.i.g.. is admittedV to the heater 15 ,unde'ra control techniquedescribed Ihereinafter to indirectly heat the first reaction products toa tempora-l -ture of 337. In addition, a heated hydrogen-rich gaseousmixture is injected into ,the products in conduit 20 upstream Vbut closeto the `chamber 18. rich stream is part of that drawn off in line8fromthe total circulating gas (recycle and make-up gases) in conduit 6.The gas in pipe 8 flows through the heat exchange-r 24, where itstempe-rature is raised to 380, and finally into gas-fired heater 26.Firing of v this heater is controlled in a unique manner which isdescribed'later; and it provides an effluent lleavingfinv conduit 28 -ata temperature of 645 which is divided by means of the three-way valve 30with 20% of the rtotal circulating gas t being introduced into pipey 32and -the lremaining' 30% passing through conduit 34 to join the rstreaction effluent in line 20. This fu-rther heating of the productstream in line 20 of course results inmore vaporization and jvaporization is completed to the desired extent'in'vapo-V, f yrizer 18.The latter is a vesselof enlarged cross section with an internaldiameter of 4.5 feet and a height of 12.5 feet which provides'favorableconditions for the substantially complete separation of the Vgaseousphase fromV the.A

liquid phase in a mixture thereof at a temperature of1360 and pressureof 695 p.s.i.g.

Based on the vrate ofy feeding pyrolysisv gasoline, 4%

of Vsaid liquid feed and reaction products thereof vaporizes inreactor12,:much more is evaporated during passage 1 through line 14 and heater15,1further substantial vaporization is produced bythe hot gas Vinjectedfrom vpipe 34 i` and only 8.5% is collected in .the liquid :phase inthe.`

vaporizer pot '18 inV addition to they circulating flux.

The gaseous phase going'overhead passes throughv the demister blanket orpad 36 of coarse steel wool designedV to catch any entrained dropletsofliquid. No substantial deposition of polymers or gums occurs in theline 14 and., 20 or heater 15, but the .liquid in the bottom of pot 18 icontains an amount of dissolved'polymer; (ASTM gum content=76 mg./ 100ml.) which issmallV but sutlcienty to foul and thereby deactivate acontactr catalyst within a fairly, short'time at desulfurizingtemperatures. A portion of the flux is continually beingremoved ataconstant rate of 200 b./d. as spent ux through the bottom line 38 underthe regulation of the flow controller 40Y This hydrogenoperating theautomatic valve 42. The rate of Withdrawing spent flux from the systemis manually reset from time to time to the minimum rate that will holdthe gum content thereof below about 100 milligrams per 100 18 The tiringof the furnace 26 for heating hydrogen-rich circulating gas iscontrolled by the automatic valve 64 operating in the fuel gas supplyline 66 in response to two temperature controllers. Temperaturecontroller 68 mls. The spent liquid ux is transferred to arerun tower 5Senses the temperature in the outlet line 28 from the (not shown).heater and maintains a temperature 645 at this point,

While an extraneous flux, may be alternatively supbut this device isreset to other temperatures as may be plied to the system at a constantrate through the line 50 required in response to the temperaturecontroller 70 connected to vaporizer 18, a suitable flux is obtainedwhich is connected to conduit 56 and maintains a temfrOm the effluent fthe iirst reactor by temporarily op- 10 perature of 515 in the chargeentering the second reactor. erating heater 15 in the manner describedhereinbefore Tb@ SeCOHd reaCiOr 58 COIltainS a bed 0f a COmpOSlte toaccumulate sufficient liquid in vaporizer pot 18 for Of Cobalt andmolybdenum suldes on a gamma alumina recycling as a flux; and thereafternormal operating con- 0f L9/10 inch Particle Sile Prepared bybydfogesuliide dtiOnS are employed in the vaporizing systeml Thetreatment in the manner described hereinbefore with a overhead or vaporphase passes through heat exchanger l Sulfur Content 0f 46% ai Operatingequilibrium and a 52 0n its Way from vaporizing chamber 18 Via ConduitWeight IalllO Of A12O3ZMOICO Of 84.717.922] I'eSpClVely. 54 t0 join thehydrogenqich gas from pipe 32 in line 56 In the initial reaction eluentthe remaining less reas the Charge for the second Stage reactor 58 Inthis active diolelins are saturated in the second reactor along passage,the heat exchanger 52 raises the temperature of With all of themono-olefin in said eliiuent in a nondethe overhead e-iuent to 485 andadmixmfe with the 20 structive manner with no substantial saturation ofarohydrogen-rich gas at about 645 further raises the temmatic compounds.The reaction conditions in the secperature of the total charge to 515 atthe reactor inlet. 0nd Stage reileiof are 3S fOiiOWS As indicatedpreviously, two unique temperature con- Inlet mpemture 515' troltechniques are employed for heating and thereby Outlet temperature 535vaporizing liquid effluent from the rst reactor to pi'e- 25 Totalpressure 685 p Si'g pare a vapor phase charge for the second reactor.First, H2 partial pressure 335 psm: the rate of il-ow of heating steam.through conduit 59 to Total H2 charged 3350 S c f b heater iscontrolled by automatic valve 60 in response Space Velocity calculatedto an external liquid level controller 62 which .is conon liquid feed 17 v /hL/vnected 1n conventional manner to sense the liquid level 30catalysts activity indexes: in separator pot 18. Since the rates ofcirculation of Desulfurization 93400, iiux liquid and removal of thespent flux are customarily Polymerization 43, held constant, a rise inthe level of liquid in pot 18 indicates that the liquid feed stock andits liquid products lrom the I nlet and Outlet tempfratures gwen I t 1SaP' are being vaporized at a lower rate. This is corrected palm. thatSignificant hydrqgenatmn ractlons Wlth Sub` automatically by the levelcontroller 62 generating a Stalmd exotherms are mkmg .place m bothreactqrs' This is borne out by a comparison of the unsaturation functionor signal in response to which valve autoindexes of the reactor chargesand etlluents of column 4 mil/cally opens to admlt more Steam Intoheater 15 nd with column 5 in Table I and also column 6 with 7. The thusVaporlze more of the rst reactor @gluem passing 40 latter two indicatethat aminor hydrogenation of diolefins hrough the ilef 15' Corfvesel 2 afau m hquld level is completed iri the iinal reactor along with theprincipal 111 the VaPOUZmg chamber mdlcate that a greater Pro'hydrogenation that saturates substantially all mono-olens Portion iSbeing VaPOIiZed, and this 1S Corrected by a and a substantially completehydrodesulfurization of orsignal from the level controller 62 to theautomatic valve game sulfur Compounds. Again there is no appreciable 6)which reduces the steam input to heater 15, and there- 45 polymerizationor deposition of coke and no noticeable fore results in a lower rate ofvap-orization in the liquid conversion of aromatic hydrocarbons tonaphthenes ocpassing therethrough. curs.

Table I i 2 a 4 5 6 7 8 9 i0 Stream Fresh Total First First Final Final.1. Stabi- Stabi- Gasol HrRich H2-Rich Reactor Reactor Reactor ReactorSep Oil lizer lized Feed Gas Gas Charge Effluent Charge Effluent as OtGas Liquid Flowlbsyhom 28,220 2, 520 14,020 35,230 35,230 39,080 39, 6801,860 920 25,400 Flow, CHIII.. 42 r13 45 240 19 er Hour:

Mol 961. 0 480. 5 C1 501. 2 250. 6 C2 68. 3 34. 2 Il S in .111. C? p p10.1 5.1

. l. 3 0. 7 C5 Saturates. 2.3 1.2 C5 Unsatuiates 3. 5 C@ Saturates 3.131a C@ Unsaturates 30. 0 Benzene 13. 6 277. 2 Toluene 0. 1 3. 6 Other01+ 15. 3

Total, Mo1s.[Hr 3 1, 560. 9 1, 134. 2

1 Measured under actual process conditions. 2Based on weight of originalliquid feed.

The gaseous product stream leaves the bottom of reactor 58 via conduit74 and is cooled by passing through heat exchangers 52 and 24respectively, as well as the cooler 76, on its way to the high pressureseparator '78 where the vapor phase is separated from the newlycondensed liquid at a temperature of 100 and pressure of 640 p.s.i.g.From this vessel the gaseous phase is taken overheadvin lines 80 and 82.About 15% of this gas is bled olf to the refinery fuel system throughpipe 84and the pressure regulator 86 which maintains the desiredpressure on the hydrogenation system. The rate of removal of thisseparator gas from the instant system is tabulated in column 8 of TableI. Most of the gaseous material, however, enters the line 90 wherein itmeets f with any make-up gas diverted from supply conduit 4 via valvedline 92 that also may require scrubbing to remove excessive hydrogensulfide. These gases are introduced into the lower half of thecombination washer 94 which is equipped with a lower caustic scrubbersection 96 bey neath a water washing section 98.

Fresh aqueous sodium hydroxide solution is admitted f in conduit 100 andjoins recirculating caustic soda soluf tion in the line 102 on its wayto the perforated scrubber trays over which it cascades downwardlyagainst the rising 1 gases. conduit 106 at the bottom and dividedbetween an exit line 108 for spent solution and conduit Htl-leading toThis alkaline liquid is drawn ol through the the water is collected inthe trough 118 and is not allowed i.

to descend therebelow and dilute the caustic scrubbing solution. Thegases rising countercurrently through the tower 94 at 640 p.s.i.g. losemost of their hydrogen sulfide content in being scrubbed irst byintimate contact with curtains of caustic soda solution, next they passthrough the demisting pad 120 intothe washing section where they arewashed with curtains of falling water to remove the last traces of H2Sas well as any entrained particles of the caustic soda solution and thenthrough the demisting pad 122.

The scrubbed and washed gases exit through the conduit 124 whichconnects with the valved by-pass line 126, that may be used to divertsome or all of the separator olf-gas around tower 94. The by-passconduitfis useful when the hydrogen sulfide content of the separator gasis low enough for a recycle gas.

These two pipes feed into the line 128 which leads to the knockout pot130 in which any entrained liquidris separated. From here thehydrogen-rich gas passes through conduit 132 to compressor 134 where itspressure is boosted sufliciently to circulate it'through the recycle gasline 6 and associated conduits in the manner described earlier.

Returning n-ow to the scrubber 94, it is apparent that an extremelyexible arrangement is shown for controlling the hydrogen sulfide contentof the circulating gases passed into the two reactors with the feed. Forexample, the operator can divide the gaseous product from separator 7Sbetween inlet line 90 of the caustic scrubber and the by-pass conduit126 in any desired proportions. Similarly, the make-up gas entering inconduit 4 can be introduced directly into circulating gas line 6 or partor all of it can be taken off via conduit 92 for treatment in thecaustic scrubber. Also, either or both of the rates of recirculation ofcaustic soda solutions in scrubber section 96 and the introduction offresh caustic soda thereto can be controlled to set the rate of reactionand removal of hydrogen sulfide from the gas stream passing through thetower. Y

The liquid phase withdrawn from the bottom of high pressure separator 78is treated in the stabilizing tower- 136 -at 180 p.s.i;g. after beingcarried in the conduit 138 through the pressure reducing valve 140 andheat exchanger 142 in which the temperature of the stream is raised to240 F. Attached to the 30-tray stabilizer are valved inlet lines 144 to146 =to introducethe charge selectivelyk and in any proportions onto the18th and 12th trays respectively counting from the bottomof the column.A reboiler 148'is provided to maintain the bottoms at a temperature ofaboutifSSfA F. and a stable, substantially saturated liquid product iswithdrawn as the product of the process via pipe 150 into heat exchanger142 at the rate given in column 10 of the table. This liquid, rich inaromatic hydrocarbons, is suitable for extraction processes, such asextraction with diethylene glycols, for the concentration of aromaticsby reason of its Y negligible Ycontent of diolens, olefns and sulfur. Itis essentially a mixture of'paraliinicrand ,aromatic hydrocarbons, and asharp separation can kreadily be obtained between these constituents.

An-overhead fraction is conveyed via the conduit 152 and cooler 154 inwhich cold waterreduces its temperature from295 to 125 in transit to thereflux yaccumulator 156. Liquid reuxis returnedfrom the Vbottom of thisaccumulator to theV stabilizer 13.61through line 158v and pump 160 at aVrate of 9840 lbs. per hour and-'aV gaseous by-product of the process iswithdrawn through the valved conduit 162 atfrthe rate set forth incolumn 9 of Table I for use as fuel or other suitable purposes.

Starting up the process described herein in a commercial plant isrelatively free of difhculties. Make-up gas obtained vfrom a catalyticreformer s charged. at; ambient temperature and the usual operatinglpartial pressure .of hydrogen into the initial reactor 12 and alsothrough the furnace 26 into the final reactor 58. This is vcontinueduntil the heat carried by theV gas from the furnace brings the secondreactor close to its normal operating temperature. Meanwhile, recyclegas is substituted for most of the f resh supply ofl hydrogen-rich gasafter a thorough purging of the system. Next, a typical reformatederived from naphtha and relatively free from runsaturated aliphaticcompounds is introduced as a temporary feed along with the` circulatinggas. An unusually high proportion of liquidaccumulates in the flashchamber 18 from kthe time theV reformate is; first chargedv until theychamber reaches normal operating temperature and none 1s withdrawnthroughthe spent fluxline at'first.y After the circulating ux system isallowed to fill up `withfthe llquid phase collecting in the Vaporizerchamber, liquid is ydrained olf in the spent flux line at an abnormallyhigh:

EXAMPLE 2 The process of Example 1 isrepeated using the same feed,equipment and reaction conditions except as otherwise specified herein.The `catalysts -are of identical composition with those-describedinExample .1 .except that the data tabulated hereinafter represents steadystate conditionsafter operating for only live days with fresh newcatalysts in both reactors.

Table Il v OPERATING CONDITIONS Flow rates-b/d.:

Pyrolysis liquid charge -3060 First reactor outlet 29,00 Flux oilcirculation 710 Spent ux v 74 Stabilizer feed 3090` Circulating gas(72.5% H2)-m.s.c.f./h.:

From .caustic and water scrubber 742 To first-stage (Pd) reactorA 341vTo furnace 26 336- To fuel line 84 21 To transfer line 34 218 Totransfer line 32 118 Make-up gas (76% H2) 98.8 Pressures-p.s.i.g.:

Second reactor, AT 4 Gas furnace outlet 28 630 Outlet of heater 350Vaporizer pot inlet 356 Circulating ux oil in line 16 350 Spacevelocities-LHSV:

First reactor 2.88

Second reactor 2.24

In comparison with Example 1, the temperatures listed immediately abovedemonstrate a considerably larger temperature rise in passing throughthe first reactor and a considerably smaller exotherm in the secondreactor. This means that a distinctly greater degree of hydrogenation istaking place in the iirst reactor and less in the second than is thecase in Example 1 where a catalyst of somewhat impaired activity is usedin the rst reactor.

The following properties of several streams are established by analyses:

Table III First Spent Stabilizer Characteristics Charge Reactor Flux Bottorns Effluent (Product) Gum, nig/100 ml 1. 5 53 Specific Gravity0.828 0. 830 0.837 O. 824 Unsaturation:

Bromine N o 18. 2 2. 9 Nil Dione No 9. 9 Nil Nil Sulfur, p.p.rn 39 l. 1

The results in Table III for catalysts with an age of ve days amount toreductions of 100% in Diene Number and 84% in Bromine Number for theproduct of the first reactor based on the values for the feed stock.Reductions of 85 and 55% respectively after 20 days of operation, and 80land 50% reductions respectively after a total of 75 days are obtainedin operating the initial reactor under the same reaction conditions. Itis estimated that the rst stage reactor can be operated for a period ofat least 6 months before regeneration of the catalyst is likely to berequired to maintain a diolen reduction of 50% while keeping the chargetemperature below 250 F.

The detailed examples given hereinabove are intended only to illustratethe invention. It will be apparent to those skilled in the art that manyother modications and variations may be made in the embodiments setforth in the examples without departing from the invention. Forinstance, standby units arranged in parallel With -alternate piping maybe provided for all equipment that requires periodic regeneration orcleaning. For simplicity and to provide comparative results, catalystsmanufactured in the same manner are employed in both of the detailedexamples; but the invention is not limited to such catalysts for a widevariety of other known hydrogenation and desulfurization catalysts maybe used insteadY Accordingly, the present invention is not to beconsidered as limited in any respect other than the recitals of theappended claims.

Certain features of the selective hydrogenation process disclosedhereinabove are also described, and claimed in concurrently filedapplication Serial No. 238,693 of Richard G. Graven et al.

What is claimed is:

1. In a process for the selective conversion of unstable liquids with apronounced tendency to deposit solids upon heating which includespartially converting unstable compounds in a liquid. feed into morestable substances Within a confined initial reaction zone underconversion conditions in which a substantial portion of said feed ismaintained in the liquid phase, vaporizing a substantial portion of theeffluent liquid from the initial conversion reaction by controlledheating and passing the gaseous phase derived from said. initialeflluent through a confined conversion zone while substantiallycompleting the conversion of unstable components of said feed; theimprovement which comprises separating in an enlarged separation zone aliquid flux in an amount equal to at least about 0.5% of said liquidfeed from said gaseous phase, withdrawing said liquid` flux at asubstantially constant rate from a pool thereof maintained in saidseparation zone and regulating said controlled heating operation indirect response to the rate of collecting said liquid ux in saidseparation zone as determined from the level of said. pool.

2. The process according to claim 1 in which said controlled heating iseffected during passage of said initial eflluent through a restrictedtransfer conduit.

3. A process according to claim 1l in which said liquid flux isseparated in an amount equal to at least about 5% of said liquid` feed.

4. A process according to claim 1 in which a substantial portion of saidliquid flux is initial effluent liquid.

5. A process according to claim 1 in which a major portion of saidliquid iiux is a material which has been subjected at least to saidinitial conversion reaction.

6. A process according to claim 1 in which a substantially constantportion of said initial effluent is retained in the liquid. phase andseparated from the gaseous phase thereof in said separation zone.

7. A process according to claim 1 in which the breadth of the boilingrange of said liquid feed does not exceed about 150 on the Fahrenheitscale.

8. A process according to claim 1 in which the breath of the boilingrange of said liquid feed does not exceed about on the Fahrenheit scale.

9. A process according to claim 1 in which said initial eilluent liquidis separated in said enlarged separation zone in an amount equal tobetween about 1 and 10% of said liquid feed.

10. A process according to claim 1l in which an indirect heat exchangeris employed in said controlled heating operation.

11. A process according to claim 1 in which a stable reactant in gaseousform is introduced into said transfer conduit immediately upstream ofsaid. separation zone at a substantially higher temperature than that ofsaid heated initial effluent.

12. A process according to claim 1 in which an amount of said initialeffluent equal to between about 20 and 50% of said liquid feed isretained in the liquid phase in said controlled. heating operation andfurther vaporization of said heated initial effluent is produced by theinjection therein upstream of said separation Zone of a stable gaseousreactant at a temperature substantially higher than that of saidefliuent.

13. A process according to claim 1 in which a stable gaseous reactant isinjected into said gaseous phase at a location between said separationzone and said conversion Zone at an injection temperature substantiallyhigher than that of said gaseous phase, which injection temperature isregulated in response to the inlet temperature of said conversion zoneto maintain said inlet temperature constant.

14. A process according to claim 1 in which a hot stable gaseousreactant is injected into said heated initial eflluent and into saidgaseous phase both upstream and downstream, respectively, from saidseparation zone.

15. A process for the selective nondestructive hydrogenation of a liquidhydrocarbon feed boiling below about 500 F. and containingk aromatichydrocarbons, and oleiins, dioleiins and sulfur compounds whichcomprises passing said feed substantially in the liquid .phase andhydrogen through an initial hydrogenation zone-in contact with a poroussolid hydrogenation catalyst of,

high hydrogenation activity' and low polymerization activity whilecontrolling hydrogenating conditions in said zone including hydrogenpartial pressure within therange of about 200-800 p.s.i.g., hourly spacevelocity Within the range of about 0.2-l5.0 based on the volume ofliquid feed, the hydrogen charge within the range ofy about 50045000s.c.f.b. of liquid feed and feed temperature within the broad range ofabout 75-300 F., said s` fluent by controlled heating and separation ofthe gaseous phase of said hydrogenation eiiuent in an enlarged sepyaration zone in the presence of a liquid flux in an amount equal to atleast about 5% of said liquid feed, withdrawing said liquid llux from apool thereof maintained in said separation zone at a substantiallyconstant rate equal to at least about-0.5% of said liquid feed,regulating said controlled heating operation in direct response to theratey of collecting said liquid uX in said separation zone as determinedfrom the level of said pool, passing hydrogen together with gaseousmaterial derived from said vaporization step through a subsequentconversion zone in contact Withfa porous solid conversion catalyst of atleast moderate hydrogenation activity and high desulfurization activityat a substantially higher average temperature than in said initial zoneWhile controlling conversion conditions in said conversion zone,including hydrogen partial'pressure Within the range of about 200-800p.si.g.,` hourly space velocity within the range of about 0.2-6.0 basedon the volume of said liquid feed, the total hydrogen charge Within therange of about SOO-10,000 s.c.f.b.

of said liquid feed and inlet temperature Within the wide` rangeof about350-700v F., said conversion conditions being regulated to produce aconversion eluent from VWithin the range of about -275. F.

Y 24; said conversion zone with anormally liquid fraction having aBromine Number less than about 4 and an organic sulfur Vcontentbelow.20-ppm.

16. A process according yto claim "15 1in which said liquidjfeed isrichiin aromatic hydrocarbons;` and boils 17.' A process according toclaim 15 in which said initialqcatalyst. has a hydrogenationVVV activityindex of at least about 40 and, a polymerization activity index lessthan about 35 and said conversion catalyst hasV a desulfurizationactivity index of at least about =80.

18. A, process according to Vclaim` 15, in which said vaporizing step iseffected in the preseucef of 'a quantity tfnf liiquid ilux equal toatlerastV about 10% of said liquid 19; A process according ito claim y15Vin which saidV vaporizing step is effected in the presence of a liquidux l recirculating at a rateequal-to at ,least about'l0% of that of saidliquid feed .and saidfiux comprises the less ing said initial zone andsaid separation zone and a minor portion of said flux is removed atleast` intermittently from said recirculating flux to minimize the'accumula-` tion of polymeric material in said flux. Y

21. A process according to claim.20 in whichlfurther` vaporization offsaid hydrogenation eluent is produced by the injection therein upstreamof said separation zone of hydrogen at a temperature substantiallyhigher than that of said effluent.

References Cited by the Examiner UNITED STATES PATENTSv 2,952,625,9/1960 Kelley et a1., v 208-254 3,051,647 8/1962 White 20S- 2553,075,917 l/l963 Kronig et al. 208-255 3,108,947 10/1963 Stijntjes 20S-255` DELBERT E. GANTZ, Primary Examiner.

ALPHONSO D. SULLIVAN, IExaminer. v

1. IN A PROCESS FOR THE SELECTEOVE CONVERSION OF UNSTABLE LIQUIDS WITH APRONOUNCED TENDENCY TO DEPOSIT SOLIDS UPON HEATING WHICH INCLUDESPARTIALLY CONVERTING UNSTABLE COMPOUNDS IN A LIQUID FEED INTO MORESTABLE SUBSTANCES WITHIN A CONFINED INITIAL REACTION ZONE UNDERCONVERSION CONDITIONS IN WHICH A SUBSTANTIAL PORTION OF SAID FEED ISMAINTAINED IN THE LIQUID PHASE, VAPORIZING A SUBSTANTIAL PORTION OF THEEFFLUENT LIQUID FROM THE INITIAL CONVERSION REACTION BY CONTROLLEDHEATING AND PASSING THE GASEOUS PHASE DERIVED FROM SAID INITIAL EFFLUENTTHROUGH A CONFINED CONVERSION ZONE WHILE SUBSTANTIALLY COMPLETING THECONVERSION ZONE WHILE SUBSTANTIALLY COMPLETING THE CONVERSION OFUNSTABLE COMPONENTS OF SAID FEED; THE IMPROVEMENT WHICH COMPRISESSEPARATING IN AN ENLARGED SEPARATION ZONE A LIQUID FLUX IN AN AMOUNTEQUAL TO AT LEAST ABOUT 0.5% OF SAID LIUQID FEED FROM SAID GASEOUSPHASE, WITHDRAWING SAID LIQUID FLUX AT A SUBSTANTIALLY CONSTANT RATEFROM A POOL THEREOF MAINTAINED IN SAID SEPARATION ZONE AND REGULATINGSAID CONTROLLED HEATING OPERATION INDIRECT RESPONSE TO THE RATE OFCOLLECTING SAID LIQUID FLUX IN SAID SEPARATIN ZONE AS DETERMINED FROMTHE LEVEL OF SAID POOL.